A process for transalkylating aromatic hydrocarbon compounds, the process comprising introducing an aromatic hydrocarbon feed stream and a sulfur source to a transalkylation zone. The feed stream contacts a catalyst in the transalkylation zone in the presence of sulfur, and produces a reaction product stream comprising benzene and xylene. The invention includes methods to control the transalkylation process.
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1. A method for controlling an aromatic transalkylation process, the method comprising:
introducing a feed stream comprising aromatic hydrocarbon compounds to a transalkylation zone;
introducing a sulfur source to the transalkylation zone;
contacting the feed stream with a sulfided catalyst comprising an aluminosilicate zeolite component having an mor framework type selected from the group consisting of UZM-14 and mordenite; an mfi molecular sieve component; an inorganic oxide binder comprising a material selected from the group consisting of alumina, silica, zirconia, titania, thoria, boria, magnesia, chromia, stannic oxide, and combinations thereof; and between about 3% and about 5% molybdenum in the transalkylation zone under transalkylation conditions including the presence of sulfur;
producing a benzene product stream;
determining a purity of the benzene product stream; and
controlling the introduction of the sulfur source in response to the purity of the benzene product stream to obtain a product having a higher purity of benzene.
2. The process of
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This application is a Division of prior U.S. application Ser. No. 12/894,209 which was filed on Sep. 30, 2010, now U.S. Pat. No. 8,242,322, the contents of which are incorporated herein by reference thereto.
The present invention generally relates to improved processes for transalkylating aromatic hydrocarbon compounds. More particularly the invention relates to aromatic transalkylation processes producing xylenes and benzene.
Xylene isomers (“xylenes”) and benzene are produced in large volumes from petroleum by the reforming of naphtha. However, neither the xylenes nor benzene are produced in sufficient volume to meet demand. Consequently, other hydrocarbons are necessarily converted to increase the yield of the xylenes and benzene via processes such as transalkylation, disproportionation, isomerization, and dealkylation. For example, toluene commonly is dealkylated to produce benzene. Alternatively, or additionally, toluene can be disproportionated to yield benzene and C8 aromatics from which the individual xylene isomers are recovered.
More recently, development has been directed at selectively transalkylating heavier aromatics, such as C9+ aromatics, with toluene and/or benzene to increase the yield of xylenes and benzene from aromatics complexes. In this regard, a variety of catalysts have been developed for these processes. For example, a wide range of zeolites, including mordenite, have been disclosed as effective transalkylation catalysts. Shaped catalysts, multiple zeolites, metal modifiers, and treatments such as steam calcination have been described as increasing the effectiveness of the catalysts.
Known catalysts are effective for producing xylenes and benzene. Specifically, catalysts having a sufficient metal function are suitable to convert heavier aromatics, such as C9+ aromatics to xylenes and benzene and provide improved catalyst stability in a transalkylation process. However, in transalkylation processes employing such catalysts, aromatic rings may become saturated or even cleaved resulting in naphthene and acyclic paraffin (non-aromatics) co-production, which can result in a loss of valuable aromatics. Also, because some of the non-aromatics have similar boiling points to benzene (benzene co-boilers), they are not readily removed to achieve a benzene product having a desired purity for commercial applications. Although the benzene co-boilers can be fractionated or extracted with a solvent, such processes are expensive and typically require additional equipment.
Accordingly, it is desirable to provide a transalkylation process that produces a high purity benzene product. In another aspect, it is desirable to provide a transalkylation process to that produces less benzene co-boilers. Other desirable features and characteristics of the present invention will become apparent from the subsequent detailed description of the invention and the appended claims.
It has been discovered that introducing sulfur into a transalkylation zone reduces the production of benzene co-boilers and/or improves the purity of the benzene fraction of the transalkylation reaction product stream. In another aspect, the invention enables control of the purity of a benzene product stream and/or control of the amount of benzene co-boilers relative to benzene in the reaction product stream or fraction thereof.
In an embodiment, the invention is a process for transalkylating aromatic hydrocarbon compounds comprising introducing a sulfur source and the aromatic hydrocarbon compounds to a transalkylation zone. Contacting the feed stream with a catalyst in the transalkylation zone under transalkylation conditions including the presence of sulfur. Producing a reaction product stream comprising benzene and xylene. The catalyst may comprise an aluminosilicate zeolite component having an MOR framework type, an MFI molecular sieve component having a Si/Al2 molar ratio of less than 80, an inorganic oxide binder, and a metal component comprising a metal selected from the group consisting of nickel, cobalt, molybdenum, tungsten, tin, germanium, lead, indium, platinum, palladium, and combinations thereof.
In another embodiment, the invention is a method for controlling an aromatic transalkylation process comprising introducing aromatic hydrocarbon compounds and a sulfur source to a transalkylation zone. Contacting the feed stream with a catalyst in the transalkylation zone under transalkylation conditions including the presence of sulfur. Producing a benzene product stream, determining a purity of the benzene product stream, and controlling the introduction of the sulfur source in response to the purity of the benzene product stream.
The aromatic hydrocarbons to be transalkylated by processes of the invention include alkylaromatic hydrocarbons of the general formula C6H(6-n)Rn, where n is an integer from 0 to 5 and R is CH3, C2H5, C3H7, or C4H9, in any combination. Non-limiting examples include: benzene, toluene, ethylbenzene, ethyltoluenes, propylbenzenes, tetramethylbenzenes, ethyl-dimethylbenzenes, diethylbenzenes, methylethylbenzenes, methylpropylbenzenes, ethylpropylbenzenes, triethylbenzenes, trimethylbenzenes, di-isopropylbenzenes, and mixtures thereof. The feed stream may comprise lower levels of ortho-xylene, meta-xylene, and para-xylene that are desired products of the process.
As used herein, the term “transalkylation” encompasses transalkylation between and among alkyl aromatics, between benzene and alkyl aromatics, and it includes dealkylation and disproportionation, e.g., of toluene to benzene and xylene. The aromatic hydrocarbons also may comprise naphthalene and other C10 and C11 aromatics. Herein, hydrocarbon molecules may be abbreviated C1, C2, C3, . . . Cn, where “n” represents the number of carbon atoms in the hydrocarbon molecule. Such abbreviations followed by a “+” is used to denote that number of carbon atoms or more per molecule, and a “−” is used to denote that number of carbon atoms or less per molecule.
Polycyclic aromatics having from 2 to 4 rings are permitted in the feed stream of the present invention. Non-limiting examples include: indanes, naphthalenes, tetralins, decalins, biphenyls, diphenyls and fluorenes. Indane is meant to define a nine carbon atom aromatic species with one ring of six carbon atoms and one ring of five carbon atoms wherein two carbon atoms are shared. Naphthalene is meant to define a ten carbon atom aromatic species with two rings of six carbon atoms wherein two carbon atoms are shared.
The aromatic hydrocarbons to be transalkylated may be introduced to the transalkylation zone in one or more feed streams. As used herein, the term “zone” can refer to one or more equipment items and/or one or more sub-zones. Equipment items may include, for example, one or more vessels, heaters, separators, exchangers, conduits, pumps, compressors, and controllers. Additionally, an equipment item can further include one or more zones or sub-zones. In embodiments having multiple feed streams, the feed streams may be introduced separately to the transalkylation zone, or two or more of the feed streams may be combined in any manner prior to passing them into the transalkylation zone.
The feed streams may be derived from one or more sources including, without limitation, catalytic reforming, pyrolysis of naphtha, distillates or other hydrocarbons to yield light olefins and heavier aromatics-rich byproducts, and catalytic or thermal cracking of heavy oils to yield products in the gasoline range. Products from pyrolysis or other cracking operations generally will be hydrotreated according to processes well known in the industry before being charged to the transalkylation zone in order to remove sulfur, olefins and other compounds which would affect product quality. Light cycle oil also may be beneficially hydrocracked to yield lighter components which can be reformed catalytically to yield the aromatics-rich feed stream. If the feed stream is catalytic reformate, the reformer preferably is operated at high severity for high aromatics yield with a low concentration of non-aromatics in the product. The reformate may also be subjected to olefin saturation to remove potential product contaminants and materials that could polymerize to heavy non-convertibles in a transalkylation process. Such processing steps are described in U.S. Pat. No. 6,740,788 B1, which is incorporated herein by reference thereto.
A feed stream can include a substantially pure alkylaromatic hydrocarbon of from about 6 to about 15 carbon atoms, a mixture of such alkylaromatic hydrocarbons, or a hydrocarbon fraction rich in said alkylaromatics. A feed stream also may contain lesser concentrations of non-aromatics such as pentanes, hexanes, heptanes and heavier paraffins along with paraffins along with methylcyclopentane, cyclohexane and heavier naphthenes; pentanes and lighter paraffins generally will have been removed before processing. The combined transalkylation feed preferably contains no more than about 10 wt % non-aromatics; and olefins preferably are restricted to a Bromine Index of no more than about 1000, and preferably no more than about 500.
In an embodiment, at least two feed streams are introduced to the transalkylation zone, a light feed stream and a heavy feed stream. The light aromatic feed stream may comprise at least one of benzene and toluene. Preferred components of the heavy aromatic feed are C9+ aromatics, thereby effecting transalkylation of toluene and C9+ aromatics to yield additional xylenes. Benzene may also be transalkylated to yield additional toluene. Indane may be present in the heavy aromatics feed stream although it is not a desirable component to effect high yields of xylenes in the transalkylation zone effluent. C10+ aromatics also may be present, preferably in an amount of 30% or less of the heavy aromatic feed. The heavy aromatic feed stream preferably comprises at least about 90 mass % aromatics, and may be derived from the same or different known refinery and petrochemical processes as the benzene/toluene feed stream and/or may be recycled from the separation of the transalkylation effluent.
The aromatic feed to a transalkylation reaction zone is usually first heated by indirect heat exchange against the reaction product stream and then is heated to reaction temperature by exchange with a warmer stream, steam or a furnace. The feed is preferably transalkylated in the vapor phase and in the presence of hydrogen. In an embodiment a hydrogen stream is introduced to the transalkylation zone. The hydrogen stream may comprise other compounds, e.g. C1 to C4 hydrocarbons, in addition to hydrogen. Hydrogen and hydrocarbons may be recycled in the process as described below. If present, free hydrogen is associated with the feedstock and recycled hydrocarbons, if any, in an amount from about 0.1 moles per mole of aromatics up to 10 moles per mole of aromatics. This ratio of hydrogen to aromatics is also referred to as hydrogen to hydrocarbon ratio.
The feed then is passed through one or more reactors containing the transalkylation catalyst to produce a reaction product stream comprising unconverted feed and product hydrocarbons including xylenes and benzene. This reaction product stream is normally cooled by indirect heat exchange against the aromatic feed stream entering the transalkylation zone and may be further cooled through the use of air or cooling water. The reaction product stream may be separated e.g. in a vapor-liquid separator to produce a vapor phase hydrogen stream and a liquid phase reaction product stream. The vapor phase hydrogen stream includes hydrogen and light hydrocarbons which may be recycled and combined with the feed as described above. The liquid phase reaction product stream may be passed into a stripping column in which substantially all C5 and lighter hydrocarbons present are concentrated into an overhead stream and removed from the process. As used herein, the term “substantially all” means an amount generally of at least 90%, preferably at least 95%, and optimally at least 99%, by weight, of a compound or class of compounds in a stream. The stripping column also produces a net stripper bottoms stream, which is referred to herein as the transalkylation zone effluent.
The transalkylation zone effluent may be further separated in a distillation zone comprising at least one distillation column to produce a benzene product stream. Various flow schemes and combinations of distillation columns to separate transalkylation zone effluent via fractional distillation are well known in the art. In addition to the benzene product stream, the distillation zone may produce a toluene product stream, and a C8+ product stream. See, e.g., U.S. Pat. No. 7,605,295. It is also known that the transalkylation zone stripper column may be designed and operated to produce a benzene product stream. See, e.g., U.S. Pat. No. 6,740,788. Thus, the reaction product stream contains a benzene fraction that may be separated by fractional distillation to produce a benzene product stream.
In another embodiment, the transalkylation effluent is separated into a light recycle stream, a mixed C8 aromatics product, and a heavy aromatic product stream in a distillation zone. The mixed C8 aromatic product may be sent for recovery of para-xylene and/or other isomers. The light recycle stream may be diverted to other uses such as benzene and toluene recovery, but may be recycled, in part, to the transalkylation zone. The heavy recycle stream contains substantially all of the C9 and heavier aromatics and may be partially or totally recycled to the transalkylation reaction zone.
The invention also requires introducing a sulfur source into the transalkylation zone. The sulfur source may be introduced as a separate stream into the transalkylation zone, and/or the sulfur source may be combined with an aromatic feed stream and/or a hydrogen containing stream. The sulfur may be introduced as a sulfur compound or sulfur containing component. The sulfur source may decompose or otherwise produce one or more components containing sulfur in the transalkylation zone. Typical examples of sulfur sources include without limitation: hydrogen disulfide, carbon disulfide, and alkylsulfides such as methylsulfide, dimethylsulfide, dimethyldisulfide, diethylsulfide and dibutylsulfide.
The invention is distinct from the known in-situ catalyst sulfiding procedure during start-up that is used for some catalysts to provide a metal-sulfur compound on the catalyst (usually converts a metal or metal oxide to a metal sulfide). In-situ catalyst sulfiding is generally conducted to add a fixed amount of sulfur to the catalyst and after the catalyst has been sulfided, sulfur addition is stopped. Sulfur, beyond that needed to sulfide the catalyst, has been considered a contaminant. See e.g. Handbook of Petroleum Refining Processes, 2nd Edition, edited by Robert A. Meyers, 1997 McGraw-Hill, page 2.60 (limiting sulfur in the feed to 1 ppm maximum). Therefore, in-situ catalyst sulfiding is limited to a short period of time at the beginning of a process cycle. In-situ catalyst sulfiding is readily accomplished in the initial 2 days on stream of a process cycle. In an embodiment, the sulfur source is introduced after 2 days on stream of a process cycle, i.e. after 48 hours of introducing feed. In another embodiment, the sulfur source is introduced after 5 days on stream of a process cycle, and the sulfur source may be introduced after 10 days on stream of a process cycle. The immediately above embodiments, do not preclude a sulfur source from being introduced before the stated days on stream.
In an embodiment, the sulfur source is introduced to the transalkylation zone in an amount to provide more than 1 ppm-wt of sulfur based upon the mass of the feed stream. In another embodiment, the sulfur source is introduced to the transalkylation zone in an amount to provide at least about 2 ppm-wt of sulfur, and may provide at least about 4 ppm-wt of sulfur, and optionally, provides at least about 6 ppm-wt of sulfur, based upon the mass of the feed stream. In an embodiment, the sulfur source is introduced to the transalkylation zone in an amount to provide from more than 1 ppm-wt of sulfur to less than 150 ppm-wt of sulfur based upon the mass of the feed stream. In another embodiment, the sulfur source is introduced to the transalkylation zone in an amount to provide from more than 1 ppm-wt to about 125 ppm-wt, and may provide from more than 1 ppm-wt to about 100 ppm-wt, and optionally, provides from about 2 ppm-wt to about 80 ppm-wt, of sulfur based upon the mass of the feed stream. In a further embodiment, the sulfur source is introduced to the transalkylation zone in an amount to provide from about 2 ppm-wt to less than 150 ppm-wt, and may provide from about 2 ppm-wt to about 125 ppm-wt, and optionally, provides from about 2 ppm-wt to about 100 ppm-wt, of sulfur based upon the mass of the feed stream. The sulfur source may be introduced to the transalkylation zone in an amount to provide from about 2 ppm-wt to about 45 ppm-wt, and optionally, provides from about 6 ppm-wt to about 45 ppm-wt, of sulfur based upon the mass of the feed stream.
The aromatic feed stream introduced into the transalkylation zone contacts the catalyst at transalkylation conditions to produce the reaction product stream which comprises comprising unconverted feed, xylenes, and benzene. The transalkylation conditions include the presence of sulfur. That is, the sulfur source and/or a sulfur containing compound derived from the sulfur source also contacts the catalyst. The sulfur containing compound may be obtained by reaction of the sulfur source with a component in the feed and/or the hydrogen recycle stream. In an embodiment, the presence of sulfur is in the vapor phase. Without wishing to be bound by any particular theory, it is believed that the sulfur modifies or attenuates the metal function of the catalyst to reduce the amount of benzene co-boilers that would otherwise be produced, thus increasing the purity of the benzene fraction of the reaction product stream. As used herein the term “benzene co-boilers” means non-aromatic hydrocarbon compounds having 6 or 7 carbon atoms per molecule. Benzene co-boilers that are particularly difficult to separate from benzene by fractional distillation include cyclohexane; methylcyclopentane; 2,3-dimethylpentane; 3-methylhexane; and dimethylcyclopentane.
The sulfur source may be introduced into the transalkylation zone continuously at a constant rate, or the rate of sulfur source addition may vary during the process cycle, e.g. as the feed quality varies, as the catalyst deactivates, and as processing objectives such as the degree of conversion or benzene purity specification change. A process cycle runs from the initial introduction of feed until the process is discontinued to regenerate or replace the catalyst. A process cycle may be measured in a variety of ways including: time on stream (i.e. time feed is being introduced); the quantity of feed that has been introduced, e.g. mass or volume of feed processed; and the quantity of feed (mass or volume) per quantity of catalyst (mass or volume), e.g. barrels (feed) per pound (catalyst) (BPP), cubic meters of feed (e.g. at standard or normal conditions) per mass or volume of catalyst. A process cycle may be temporarily suspended or halted by discontinuing feed introduction. If suspended for an extended period of time, the unit may be held in a stand-by mode. However, the length of the process cycle will continue to accrue once feed introduction resumes, provided the catalyst has not been regenerated or replaced. In another embodiment, the sulfur source is added intermittently during the process cycle.
In a further embodiment, the sulfur source is introduced when the feed is initially introduced to the transalkylation zone or after the catalyst has been sulfided in-situ. During the initial portion of the process cycle, the catalyst usually exhibits the highest activity; therefore, the operating temperatures are usually at a minimum. These conditions typically result in higher amounts of benzene co-boilers and lower benzene purity. The catalyst ages and deactivates as the process cycle continues and to maintain the desired level of conversion the reactor temperature may be increased. The amount of sulfur source addition may be decreased as the process cycle continues and the reaction temperature increases. Without wishing to be bound by any particular theory, it is believed that benzene co-boilers are cracked more efficiently by the acid sites on the catalyst, especially the MFI zeolite, at higher temperatures. At some point in the process cycle, sulfur source introduction may no longer be required, e.g. to obtain the desired benzene product purity.
In an embodiment, the sulfur source is introduced intermittently or continuously during the initial one-third of the process cycle. In another embodiment, the sulfur source is introduced intermittently or continuously during the initial one-fifth (20%) of the process cycle; and the sulfur source may be introduced intermittently or continuously during the initial one-tenth (10%) of the process cycle. To determine fraction of a process cycle herein, the process cycle and the fraction thereof are determined on the basis of mass of feed per mass of catalyst, e.g. metric ton of feed per metric ton of catalyst. In an embodiment, the sulfur source is introduced intermittently or continuously after the initial 1% of the process cycle; and the sulfur source may be introduced intermittently or continuously after initial 2% of the process cycle. In an embodiment, the sulfur source is introduced intermittently or continuously from after the initial 1% of the process cycle to about 10% of the process cycle; and the sulfur source may be introduced intermittently or continuously from about 2% of the process cycle to about 10% of the process cycle; optionally from about 2% of the process cycle to about 15% of the process cycle.
In an embodiment, the purity of the benzene product may be controlled by adjusting the rate of sulfur introduced in response to a determination of the benzene product purity. Benzene purity may be determined by analysis of the benzene product stream and/or from a determination of the amount of benzene co-boilers relative to the amount of benzene. For example, the amount of benzene co-boilers relative to the amount of benzene may be determined in any of the reaction product stream, the transalkylation zone effluent stream, and the benzene product stream. The determination of benzene purity and the relative amount of the benzene co-boilers may be performed by manually obtaining samples and analyzing them off-line or by automatic on-line analysis. Changing the rate of sulfur source introduction in response to the determination may be done manually or through the use of a control system as is known in the art. The manual and automatic modes of these steps may be combined in any manner. For example, an on-line analyzer may determine the amount of benzene co-boilers relative to the amount of benzene in the reaction product stream and send a signal to a process controller. The process controller may in turn send a signal to a control valve that regulates the rate of sulfur addition. Algorithms, such as to convert the relative amount of co-boilers to a purity of the benzene product steam may include terms to account for the fractional distillation efficiency and/or time lags in the process, and may be applied in or by any of the analyzer, controller, and control valve to generate or interpret the signals. Such techniques are well known in the art of process control.
Contacting the feed and sulfur with the catalyst can be effected in any conventional or otherwise convenient manner and may occur as a batch or continuous type of operation. In an embodiment, the catalyst is disposed in one or more fixed beds in a reaction zone of a vertical reactor with the aromatic feed and sulfur charged through the bed in an upflow or downflow manner. Transalkylation conditions may include a temperature in a range of from about 200° C. to about 540° C., preferably between about 200° C. to about 480° C.; a pressure in a range of from about 100 kPa to about 6 MPa absolute; and a weight hourly space velocity (WHSV, i.e., weight of aromatic feed introduced per weight of catalyst per hour) in a range of from about 0.1 to about 20 hr-1.
In an embodiment, the invention includes a transalkylation catalyst comprising: an aluminosilicate zeolite having an MOR framework type, an MFI molecular sieve having a Si/Al2 molar ratio of less than 80, a metal component comprising a metal selected from the group consisting of nickel, cobalt, molybdenum, tungsten, tin, germanium, lead, indium, platinum, palladium, and combinations thereof, and an inorganic oxide binder.
Aluminosilicate zeolite having an MOR framework is described in Atlas of Zeolite Framework Types, 6th Revised Edition, C. H. Baerlocher, L. B. McCusker, and D. H. Olson, editors, Elsevier (2007), pp. 218-219. The MOR framework comprises four- and five-membered rings of SiO4 and AlO4 tetrahedra to form a crystal lattice comprising 12-ring channels running parallel along a crystal axis to give a tubular configuration. In an embodiment, the aluminosilicate zeolite having an MOR framework comprises mordenite. Where mordenite is a component of the catalyst, the mordenite preferably has a Si/Al2 molar ratio of less than about 40. The Si/Al2 molar ratio of mordenite in an embodiment is less than about 25, and in another embodiment the mordenite Si/Al2 molar ratio is between about 15 and about 25. Mordenite may be synthesized with a Si/Al2 molar ratio of between about 10 and about 20. Mordenite is preferably at least partially in the hydrogen form and/or may be dealuminated by a variety of techniques, e.g. steaming, and acid extraction of aluminum to increase the Si/Al2 ratio of the mordenite.
In another embodiment, the aluminosilicate zeolite having an MOR framework comprises UZM-14. UZM-14 is described in U.S. Pat. No. 7,687,423, which is incorporated herein by reference in its entirety. UZM-14 comprises globular aggregates of crystallites having a MOR framework type comprising 12-ring channels, and one or more of the following distinctive characteristics: a mesopore volume of at least about 0.10 cc/gram, preferably at least about 0.13 cc/gram, more preferably at least about 0.2 cc/gram; a mean crystallite length parallel to the direction of the 12-ring channels of about 60 nm or less, preferably about 50 nm or less; a Si/Al2 mole ratio of between about 8 and about 50, and preferably is no more than about 30; and at least about 1×1019 12-ring channel openings per gram of UZM-14 material.
In an embodiment, UZM-14 comprises globular aggregates of crystallites having a MOR framework type comprising 12-ring channels, a silica-alumina mole ratio of from about 8 to no more than about 30, a mesopore volume of at least about 0.10 cc/gram, and a mean crystallite length parallel to the direction of the 12-ring channels of about 60 nm or less.
UZM-14 has an empirical composition in the as-synthesized form on an anhydrous basis expressed by the empirical formula:
Mmn+Rrp+Al1-xSiyOz
where M is at least one exchangeable cation and is selected from the group consisting of alkali and alkaline earth metals including but not limited to lithium, sodium, potassium, rubidium, cesium, calcium, strontium, barium and mixtures thereof. R is at least one organic cation selected from the group consisting of protonated amines, protonated diamines, quaternary ammonium ions, diquaternary ammonium ions, protonated alkanolamines, and quaternized alkanolammonium ions. Relating the components, “m” is the mole ratio of M to Al and varies from about 0.05 to about 0.95; “r” is the mole ratio of R to Al and has a value of about 0.05 to about 0.95; “n” is the weighted average valence of M and has a value of about 1 to about 2; “p” is the weighted average valence of R and has a value of about 1 to about 2; “y” is the mole ratio of Si to Al and varies from about 3 to about 50; and “z” is the mole ratio of 0 to Al and has a value determined by the equation: z=(m·n+r·p+3+4y)/2.
The catalyst also includes an MFI molecular sieve having a Si/Al2 molar ratio of less than 80. Zeolites having an MFI type framework are described in Atlas of Zeolite Framework Types, 6th Revised Edition, C. H. Baerlocher, L. B. McCusker, and D. H. Olson, editors, Elsevier (2007). MFI type zeolites have a 3-dimensional 10-ring channel system: [100] 10-MR 5.1×5.5 Å and [010] 10-MR 5.3×5.6 Å. In an embodiment, MFI molecular sieves used in the catalysts of this invention have a Si/Al2 molar ratio of less than about 40, preferably less than about 25, for example, between about 15 to about 25. An example of a suitable MFI molecular sieve for inclusion in the catalyst includes, but is not limited to, ZSM-5, which is disclosed in U.S. Pat. No. 3,702,886, incorporated herein, by reference thereto. Suitable MFI molecular sieves are also available, for example, from Zeolyst International of Conschocken, Pa. and Tosoh Corporation of Tokyo, Japan.
In an embodiment, the MFI molecular sieve has a “Total Acidity” of at least about 0.15, preferably at least about 0.25, and more preferably at least about 0.4, for example, 0.4 to 0.8. Total Acidity is determined by Ammonia Temperature Programmed Desorption (Ammonia TPD). The Total Acidity of the MFI molecular sieve may be that of the MFI to be used in making the catalyst of the invention or may be achieved during the preparation of the catalyst. Typically, the MFI molecular sieve is at least partially in the hydrogen form in the finished catalyst. The Ammonia TPD process involves first heating a sample (about 250 milligrams) of molecular sieve at a rate of about 5° C. per minute to a temperature of about 550° C. in the presence of a 20 volume percent oxygen in helium atmosphere (flow rate of about 100 milliliters per minute). After a hold of about one hour, helium is used to flush the system (about 15 minutes) and the sample is cooled to about 150° C. The sample is then saturated with pulses of ammonia in helium at about 40 milliliters per minute. The total amount of ammonia used is greatly in excess of the amount required to saturate all the acid sites on the sample. The sample is purged with helium (about 40 milliliters per minute) for about 8 hours to remove physically adsorbed ammonia. With the helium purge continuing, the temperature is increased at a rate of about 10° C. per minute to a final temperature of 600° C. The amount of ammonia desorbed is monitored using a calibrated thermal conductivity detector. The total amount of ammonia is found by integration. Dividing the total amount of ammonia by the dry weight of the sample yields the Total Acidity. As used herein, values of Total Acidity are given in units of millimoles of ammonia per gram of dry sample.
The inorganic oxide binder of the catalyst comprises such materials as alumina, silica, zirconia, titania, thoria, boria, magnesia, chromia, stannic oxide, and the like as well as combinations and composites thereof, for example silica-alumina, alumina-zirconia, alumina-titania, aluminum phosphate, and the like. Alumina is a preferred refractory inorganic oxide binder. As is well known in the art, a precursor of the desired refractory inorganic oxide may be used to form, bind, and/or otherwise prepare the catalyst. Such binder precursors or sources may be converted into a refractory inorganic oxide binder, e.g. by calcination. The alumina may be any of the various aluminum oxides, hydroxides, and gels, including boehmite, pseudo-boehmite, gibbsite, bayerite, and the like, especially transition and gamma aluminas. Suitable aluminas are commercially available, e.g. under the trade names CATAPAL B and VERSAL 250.
The metal component of the catalyst comprises a metal selected from the group consisting of nickel, cobalt, molybdenum, tungsten, tin, germanium, lead, indium, platinum, palladium, and combinations thereof. In an embodiment, metal component comprises a metal selected from the group consisting of nickel, cobalt, molybdenum, tungsten, platinum, palladium, and combinations thereof. In an embodiment the metal content of the catalyst ranges from about 0.01 wt % to about 10.0 wt % as the metal based upon the total weight of the catalyst.
The metal component may be incorporated into the catalyst in any suitable manner such as comulling, coprecipitation or cogellation with the carrier material, ion exchange, or impregnation. The metal component may exist within the final catalyst as a compound such as an oxide, sulfide, halide, or oxyhalide, in chemical combination with one or more of the other ingredients of the composite, or as an elemental metal. One method of preparing the catalyst involves the use of a water-soluble or solvent-soluble, decomposable compound of the metal to impregnate the molecular sieve-containing support. Alternatively, a metal compound may be added at the time of compositing the molecular sieve component and binder.
The weight ratio of the MFI molecular sieve component to the aluminosilicate zeolite having the MOR framework may range from about 1:10 to 5:1, preferably from about 1:10 to 2:1. In an embodiment, the aluminosilicate zeolite component having the MOR framework comprises from about 20 wt % to about 80 wt % of the catalyst, the MFI molecular sieve component comprises from about 10 wt % to about 70 wt % of the catalyst, and the inorganic oxide binder comprises between about 1 wt % and about 40 wt % of the catalyst.
The catalyst may optionally include an additional molecular sieve component preferably selected from one or more of MEL, EUO, FER, MFS, MTT, MTW, MWW, MAZ, TON and FAU (IUPAC Commission on Zeolite Nomenclature) and UZM-8 (see U.S. Pat. No. 6,756,030 which is herein incorporated by reference in its entirety). The catalyst may optionally include a fluoride component in an amount ranging from about 0.1 wt % to about 5.0 wt % of fluoride based upon the total weight of the catalyst. The fluoride component may be incorporated into the catalyst by any known technique, e.g. impregnation.
The techniques used to prepare the catalyst are well known to those of ordinary skill in the art. The catalyst can be formed by combining the aluminosilicate zeolite component having the MOR framework, the MFI molecular sieve component, and the inorganic oxide binder and/or a precursor thereof in any conventional or otherwise convenient manner to form spheres, pills, pellets, granules, extrudates, or other suitable particle shapes. For example, finely divided aluminosilicate zeolite having the MOR framework and MFI molecular sieve particles, and metal salt particles can be dispersed in an alumina sol, and the mixture in turn dispersed as droplets in a hot oil bath whereby gelation occurs with the formation of spheroidal gel particles. A preferred method comprises mixing a finely divided form of the selected aluminosilicate zeolite having the MOR framework, MFI molecular sieve particles, a binder and/or precursor thereof, with a metal salt and, optionally, a lubricant; and compressing the mixture into pills or pellets. Alternatively, and still more preferably, the aluminosilicate zeolite having the MOR framework, MFI molecular sieve particles, binder and/or precursor thereof, and metal salt are combined and admixed with a peptizing agent in a mixer-muller, a dilute nitric acid being one example of the suitable peptizing agent. The resulting dough can be pressured through a die or orifice of predetermined size to form extrudate particles which can be dried and calcined and utilized as such. A multitude of different extrudate shapes are possible, including, but not limited to, cylinders, cloverleaf, dumbbell and symmetrical and asymmetrical polylobes, with a trilobe form being favored. The extrudates also may be formed into spheres by means of a spinning disc or drum. The variously formed particles are then usually dried and/or calcined.
If the metal component is not included in the above forming steps, or if an additional metal component is to be included, the formed particles produced above can be impregnated with a soluble, decomposable compound containing the metal component to form a composite. For example, when the metal component comprises molybdenum, typical compounds which may be employed include ammonium heptamolybdate, alkali metal molybdates (also peroxo-, di-, tri-, tetra-, hepta-, octa-, or tetradecamolybdate), molybdic acid, phosphomolybdic acid, Mo—P heteropolyanion compounds, acetyl acetonates, Mo(0) metal, Mo oxides, Mo peroxo complexes, and mixtures thereof. The composite is preferably calcined in an air atmosphere at a temperature of from about 425° C. to about 750° C., preferably at a temperature of from about 475° C. to about 600° C., over a period of from about 0.5 to about 10 hours. Typically, the formed particles are also calcined at similar conditions prior to the impregnation step. The catalyst preparation may include various optional steps such as drying and steaming which are well known in the art.
A catalyst was prepared by blending of UZM-14, MFI zeolite, and Catapal B to obtain a 50% UZM-14, 25% MFI zeolite, and 25% Catapal B mixture on a volatile free weight basis. The mixture also included a solution of ammonium heptamolybdate to obtain 5 wt % molybdenum in the final catalyst, and a solution of diluted nitric acid as the peptizing agent to form a dough. The dough was extruded as a trilobe and the catalyst was calcined at 500° C. for 2 hours in dry air. The UZM-14 zeolite used in this example was prepared according to Example 1 of U.S. Pat. No. 7,687,423 except the crystallization temperature was 140° C. and commercial scale equipment was used. The as synthesized, washed, and dried UZM-14 zeolite had the following properties: a SiO2/Al2O3 molar ratio of 17.7, a BET surface area of 409 m2/g, a Langmuir surface area of 637 m2/g, a micropore volume of 0.21 cc/gram, and a mesopore volume of 0.24 cc/gram. Prior to incorporation in the catalyst, the UZM-14 was ion exchanged with an ammonium sulfate solution to reduce the sodium content of the zeolite to less than 0.05 wt % sodium. The MFI zeolite was CBV 2314, a ZSM-5 material with SiO2/Al2O3 of 23, obtained from Zeolyst International. Catapal B was purchased from Vista Chemical Company. The finished catalyst had a BET surface area of 362 m2/g, a total pore volume of 0.38 cc/g and a piece density of 1.097 g/cc.
The catalyst in this example was prepared in the same way as described in example 1, except that it was extruded as a cylinder and after extrusion the catalyst was calcined at 540° C. for 2 hours in air with 7.5 mole % steam. This catalyst had a BET surface area of 361 m2/g, a total pore volume of 0.39 cc/g, and a piece density of 1.120 g/cc.
The catalysts of Examples 1 and 2 were tested in an aromatics transalkylation test following a standard transalkylation testing procedure. Prior to testing, the catalysts were sulfided in-situ as is well known in the art to convert the MoO3 of the calcined catalyst, at least partially, to molybdenum sulfide. An objective of catalyst sulfiding is to add a fixed amount of sulfur to the catalyst. This was accomplished by passing excess dimethyl disulfide (DMDS), equivalent to 150 ppm-wt as sulfur, in the feed over the catalyst at a temperature of 280° C., a pressure of 250 psig, a weight hourly space velocity of 4 and a hydrogen to hydrocarbon ratio of 6 for 26 hours. Catalyst sulfiding was continued at a temperature of 360° C. for 12 hours, followed by 2 hours at 350° C. This procedure provides a sulfided catalyst with a relatively fixed sulfur content such that longer sulfiding with excess DMDS will not increase the sulfur content of the catalyst any further. After the sulfiding procedure was complete, feed without a sulfur source was introduced to the catalyst for 10 hours. Then, standard transalkylation testing was conducted without sulfur at temperatures ranging from 350° C. to 405° C. for a total time of about 70 hours.
Thereafter, the overall conversion, calculated as the net disappearance of toluene, C9 and C10 aromatics from the feed stream, i.e. (mass of toluene, C9 and C10 aromatics in the feed minus mass of toluene, C9 and C10 aromatics in the reaction product)/mass of toluene, C9 and C10 aromatics in the feed, on a weight percent basis; and the benzene purity, calculated as benzene/(benzene plus C6 and C7 non-aromatics) on a weight percent basis, were measured at 365° C., a pressure of 250 psig, a weight hourly space velocity of 4 and a hydrogen to hydrocarbon ratio of 6 during each of the four stages of the test. Product samples were obtained hourly throughout all four stages of the test for analysis. Stage 1 established a base line performance for the first 6 hours of testing with no sulfur addition at the above test conditions. Stage 2 then began when DMDS was added continuously to the feed to obtain 6 ppm-wt of sulfur based on the mass of the feed stream. After a transition period of about 5 hours, steady state performance was observed. For Stage 3, the rate of DMDS introduction was increased to obtain 45 ppm-wt of sulfur based on the mass of the feed stream. After a transition period of about 2 hours, steady state performance was observed. Finally, the sulfur addition as DMDS was discontinued to begin Stage 4. The feed had nominally the composition in weight percent given in Table 1. The results for the steady state operation of each stage are reported below in Table 2. No results are reported for Stage 4 because after sulfur addition was discontinued the benzene purity and overall conversion both trended back towards the sulfur free operating base line of Stage 1 showing the reversible nature of sulfur addition during the transalkylation process. The tests ended after about 20 hours after DMDS addition stopped and before steady state was attained.
TABLE 1
Toluene
75
Propylbenzene
2.0
Methylethylbenzene
10
Trimethylbenzene
9.7
Indane
0.8
Methylpropylbenzene
1.0
Diethylbenzene
0.4
Dimethylethylbenzene
1.0
C11+ aromatics
0.1
TABLE 2
Stage
Example
Benzene
Overall
(S added, ppm-wt)
#
Purity, wt %
Conversion, wt %
1 (0 ppm-wt)
1
99.3
50.0
1 (0 ppm-wt)
2
99.6
49.9
2 (6 ppm-wt)
1
99.6
48.5
2 (6 ppm-wt)
2
99.8
48.4
3 (45 ppm-wt)
1
99.7
47.5
3 (45 ppm-wt)
2
99.85
47.4
The results show that even low levels of sulfur addition to a transalkylation zone provide a significant improvement in benzene purity. The results indicate that sulfur addition below 6 ppm-wt would provide a benefit. Increasing the sulfur addition from 6 to 45 ppm-wt provided an additional benefit but not to the same degree. This is not surprising given the diminishing returns expected as the limit of 100% benzene purity is approached. Although the data shows the overall conversion declined with sulfur addition, constant conversion may be easily maintained by increasing the operating temperature. For the conditions and performance exhibited in these tests, it is estimated at constant conversion the increase in benzene purity would range from 0.3 wt % to 0.5 wt % by use of the invention. Larger benefits are expected when the baseline benzene purity (at no sulfur injection) is lower.
The observations above suggest that the sulfur source or resulting sulfur containing component interacts with the molybdenum sulfide on the catalyst. It is believed this further attenuates the metal function and reduces the amount of benzene co-boilers formed by saturation of aromatics. It is also believed to lead to less complete saturation of dealkylation products which allows more realkylation to aromatic rings and manifests itself as a lower overall dealkylation activity, which contributes to the lower overall conversion observed here.
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